Two-stage catalytic conversion process for producing naphthalene and an aromatic gasoline from cycle oils



GAS

Nov. 1, 1960 B. S. FRIEDMAN TWO-STAGE CATALYTIC CONVERSION PROCESS FORPRODUCING NAPHTHALENE AND AN AROMATIC GASOLINE FROM CYCLE OILS FiledAug. 29, 1956 1') LL! 2 2 2 9 5 1, 5 2% 5% 1 2 O r- I E 55 92 :1 25% O 2E O 2 2 2 m 2 E SEPARATOR 8% REACTOR 3 g0 95% m g n x E .J I w 9 (O a gA SEPARATOR :5

d] IN REACTOR M Q r. Id 9 J0 O INVENTOR BERNARD S. FRIEDMAN BY W44 MgUnited S at A TWO-STAGE CATALYTIC CONVERSION PROCESS FOR PRODUCINGNAPHTHALENE AND AN AROMATIC GASOLINE FROM CYCLE OILS Bernard S.Friedman, Chicago, Ill., assignor to Sinclair Rfelfi/lnlng Company, NewYork, N.Y., a corporation ame Filed Aug. 29, 1956, Ser. No. 606,822

9 Claims. (Cl. 208-60) This invention relates to the production ofcommercial grade naphthalene and, more specifically, to the product1onof commercial grade naphthalene and high octane aromatic gasoline. Thepresent invention is particularly concerned with a two-stagehydroconversion process in which a light cycle oil derived frompetroleum is converted to naphthalene, alkylated naphthalenes and highoctane number aromatic gasoline through the utilization of certainhydroforming catalysts in the presence of by drogen. This application isa continuation-in-part of my copending application Serial No. 434,059,filed June 2, 1954, now abandoned.

High octane gasoline and naphthalenes are highly desirable productswhich can be derived from petroleum hydrocarbons, and many refiners arein search of new methods of producing these materials from inexpensivecharge stocks. The present invention is concerned with a two-stageprocess for manufacturing these products through catalytic conversion oflight cycle oil stocks which are usually considered too refractory to beemployed 'as feeds in conventional catalytic cracking procedures sincethe amount of coke formed on the catalysts at the conditions necessaryfor cracking is prohibitive.

I am aware that the production of naphthalene from cycle oils has beenproposed, for instance, through utilization of the water gas reaction orhydropyrolysis. Although such processes may give desirable amounts ofnaphthalene, other valuable constituents of the charge stock areconverted to less valuable products such as gas due to the severeconditions needed to effect the dealkylation of the alkyl naphthalenesin the cycle oils. In the present invention, I have devised a methodwhich affords good yields of naphthalene from light cycle oils and atthe same time transforms non-naphthalenic constituents to valuable highoctane number aromatic gasoline.

Although my process will be defined below as a twostage hydroconversionof light cycle oil in the presence of non-carbon based hydroformingcatalysts at certain reaction conditions, these conditions can beselected or varied within the ranges disclosed to emphasize particularchemical reactions. In the first stage (StageI) of the present processthe reaction conditions can be maintained to convert the major portionof the non-naphthalenic constituents of the feed stock to; aromatic andnon-aromatic gasoline with a minor amount or substantially none of theseconstituents being aromatized to the C to C range. The gasoline is thenseparated and the remaining aromatic-rich oil containing naphthalenicderivatives is charged to the second or more severe reaction (Stage II)of the process to produce good yields of commercial grade naphthalene,alkylated naphthalenes and additional amounts of aromatic gasoline.

Alternatively, the Stage I conditions can be selected to effect whollyor predominantly an aromatization re action with little or no aromaticand non-aromatic gasoline being produced. Although it is possible toproduce no substantial amount of gasoline in the first stage which wouldobviate the desirability of a gasoline. removal 2,958,64 Patented Nov.1, 1960 operation, this stage will in most instances, produce at least asmall amount of gasoline which should be removed from the higher boilingaromatic concentrate before it is subjected to the second stagereaction. In this variation of the system among the products from thesecond stage are again naphthalene, alkylated naphthalenes and highoctane aromatic gasoline. In commercial practice, an operator can varythe first stage reaction conditions within the prescribed limits as hechooses; however, in each instance the ultimate products are essentiallythe same. It may even be desirable to select the conditions for thefirst stage reaction so that a substantial amount of gasoline is made;yet non-naphthalenes in the feed stock could also be converted insubstantial amounts to C to C aromatics.

In the initial stage the reaction is regulated in severity so thatvaluable non-naphthalenic constituents of the feed are converted togasoline with a minimum loss through production of less valuable gas. Ingeneral, the reaction temperature can vary from about 900 to l200 F.,while the pressure is maintained at least about atmospheric, forinstance atmospheric to 1000 p.s.i.g. or higher, preferably less than1000 p.s.i.g. The preferred reaction temperature is from about 900 to1050 P. which affords especially high octane gasoline as a product. Thespace velocity, that is the volume of liquid hydrocarbon per hour pervolume of catalyst (LHSV), can vary from about 0.1 to 20 LHSV; however,it must be correlated with temperature in order that neither too littlenor too much conversion is effected in the first stage reaction zone.The severity of the first stage reaction is controlled so that thearomatic oil (product boiling above 400 F.) contains at least about 40weight percent, preferably at least about 60 weight percent, ofnaphthalenic aromatics and not more than about 30 Weight percent,preferably not more than about 20' weight percent of non-aromaticcomponents. Should insufficient naphthalenic aromatics be present in thearomatic oil derived by a single pass through the first stage reactionzone, this oil could be recycled to Stage I for further conversion. Inthis stage the severity is not increased materially beyond the pointwhere more than about 10 weight percent of naphthalene is in thearomatic oil. Also it is preferred to operate the first stage reactionso that the yield of liquid product is over 70 weight percent of thefeed. The space velocity will usually be from about 0.5 to 8 LHSV,preferably about 1 to 3 LHSV. A space velocity of about 2.0 LHSV isparticularly preferred in producing high octane gasoline. Thetemperature and space velocity employed will be affected by the extentof conversion desired in the first stage which in turn is a function ofthe choice of catalyst, the hydrogen partial pressure and to some extentthe total pressure and the degree of catalyst deactivation or the amountof coke deposited on the catalyst at a given time in the processingcycle.

Sufiicient hydrogen must be present in the first stage reaction zone toeffect the reaction and maintain the activity of the catalyst throughdecreasing the amount of coke lay-down. In the first stage of theprocess the use of about 3 to 10 or 20 mols of hydrogen for each mol offeed stock produces satisfactory results. The use of about 5 to 6 molsof hydrogen per mol of feed is preferred, and the benefits derived byusing more than about 10 mols of hydrogen usually do not justify theincreased expense. of course, the specific amount of hydrogen employedwill vary with the feed stock charged and the reaction conditionsobserved. Under certain select conditions of temperature and pressure,e.g. 1000 F. and 50 p.s.i.g., operating with recycled tail gas, the netconsumption of hydrogen may be reduced so that an extraneous source ofhydrogen is not required. Also,

silica-magnesia, or silica-zirconia.

gasoline and the aromatic-rich oil.

3 when operating with recycled tail gas, hydrogen sulfide is preferablyremoved from such gases.

The catalysts which can be used in both stages of the present processare non-carbon based hydroforrningcatalysts which can contain chromium,molybdenum, tungsten, cobalt, or vanadium (or mixtures of these)preferably deposited as oxides on non-combustible carriers such asalumina, titania, thoria, zirconia, silica, silica alumina, Platinum,palladium, rhodium, and other rare metals may be employed preferably'inthe metallic state supported on the same noncombustible type ofcarriers. In operating my process on .a cyclic or continuous basis, thecatalyst can be regenerated at intervals by treatment with air or oxygenat temperatures above 900 F.

The preferred catalyst is chromia-alumina since it affords anexceptionally selective action leading to the formation in good yieldsof thehigh octane aromatic Molybdena-alumina catalysts are, for example,not as desirable as chromiaalumina, since the former may produce morecarbon and sulfur on the catalyst in both Stage I and Stage II ,of theprocess. chromia-alumina catalysts which are generally known in thepetroleum processing field can be .utilized; and these catalysts usuallycontain from about 1.0% to about 25% or more by weight of chromia. Theactivity of the chromia-alumina system can be appreciably enhanced bythe use of promoters which can in clude for instance silica, berryllium,boron, potassium and cerium. A specific catalyst found effectivecontains .about 12% chromia, 86% alumina, and about 2% magnesiumoxidejas a promoter.

The feed stocks which can be employed in the initial reaction arepetroleum light cycle oils boiling generally in the range from about 400to 650. F. Also, aromatic extracts of these stocks can be used. Thelight cycle oils usually are composed of at least about 40 weightpercent of non-aromatics and these may comprise about 50 to 65% or moreof the oil. The feed stock can be --desulfurized but this is notconsidered essential as a -desulfurization reaction is effected in eachstage. of the .present process and most of the sulfur present in thefeed 'is removed through conversion to hydrogen sulfide in the case ofchromia catalysts, or in part deposited on the catalyst as metallicsulfide, in the case of molybdenum catalysts.

In the second stage of the present process the aromatic oil from theinitial reaction is treated under more severe conditions to insure theproduction of (a) sufiicient yield of naphthalene in highconcentrationin its boiling range, -(b) conversion of benzenoid aromatics andnon-aromatic .hydrocarbons to highly aromatic gasoline, and (c) toelfect removal of nitrogenandsulfur-containing impuri- .ties from thedesired products. The temperature in the .secoud stage reaction can varyfrom about 900 to 1200 F. whilethe space velocity will generally be fromabout 0.1 to 3 LHSV. The reaction pressure and amount of :hydrogensupplied can be of the same range as specified for the initial stage ofthe process providing that one .or more of the operating conditions isconsiderably more severe than employed in Stage I. Compared with Stage Ithe desired severity of the-Stage II reaction may be effected by raisingthe temperature by at least about 125 obtained by a partial change ofmore than one of these variables in the indicated directions. It ispreferred to operate the second stage ofmy process so that the liquidproduct yield is over about 70 weight percent of the total fee d to thisstage and the product fraction boiling between 400 to 460 F. contains atleast about 40 weight percent naphthalene and at least about 50 weightpercent total aromatics.

The products of the second stage reaction include aromatic gasoline,naphthalene, beta-methylnaphthalene and other alkyl-naphthalenes. Theseproducts can be separated by fractional distillation. The naphthaleneand beta-methyl-naphthalene are of high purity, substantially free fromsulfurand nitrogen-containing impurities, and are practically completelyaromatic. Beta-methylnaphthalene and the other alkylnaphthalenes can beconverted to naphthalene as for instance by recycling to the secondstage of my process the aromatic oil boiling above about 400 F. fromwhich the naphthalene has been separated. 2,6-dimethylnaphthalene may beisolated by cooling, filtering, and recrystallization from the'dimethylnaphthalene fraction either from the first or second stageproducts. Also, when betamethylnaphthalene and/or 2,6-dirnethylnaphthalene .aswell as naphthalene are separated fromthereaction product of the second stage, the remaining oil boiling'aboveabout 400 F. can be re 'cycledto the second stage reaction.

When utilizing a chromia-alumina catalyst in the re- .action of StageI'of this invention, I prefer the severity in operating conditions in,the first stage to be equivalent .toor defined by 1000 to 1100 F., -2LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure. In the secondstage when employing thisccatalyst the preferred severity is equivalentto .or. ;definedxby 1050 to ll50 F.,

0.5 LHSV and 200to 5.00 p.s.i,.g. hydrogen partial. pres- :sure.Different catalysts do, .of course, re'quire differ- -.ent operatingconditions;;to produce the best results.

Thus, if a molybdena-alumina catalystzbe employed the preferredseveritiesinthe separate stages are, respectively, equivalent to. ,ordefinedzby-a900 to 1000 F., 2

LHSV, 50 to 400 p.s.i.g.hydrogen partial pressure and 900 to 1000 F.,1.0'LHSV,; 200 to'500 p.s.i.g. hydrogen partial pressure.

Thev process of the present invention can be effected,

for instance, in asingle reactor system or in a systemincluding'tworeactors. Intthe single reactor system the initial stagewill be effected and. the products separated as by fractionation.Thearomatic-rich oil is collected and then converted in a blocked-outoperation to naphthalene and additional amounts of aromatic gasoline inthe reactor after ithe'desiredamount of light cycle oil has beenprocessed. In a two-reactor system the separate stages of the processcan be conducted in the separate reactors. Also, in any reaction systemprovision can be made for regenerating the catalyst. 'Also, both stagesmay be conducted with fluidized or moving bed type of catalyticreactors.

It is not possible to obtain satisfactory yields of gasolineandnaphthalene by operating with one reactor continually recyclingthe 400F. plus product from which naphthalenehas been recovered. In the first-place, operatingconditions severe enough to effect demethylationoff-methylated naphthalenes are much too severe for converting thenon-aromatics to gasoline, and will produce more dry 'gas than theeconomics will per- .mit. Conversely .when'the conditions are optimumfor gasoline production, they are not severe enough to deconcurrentlywith the cycle oil directly to the initial stage 'reactorto produce theneeded hydrogen while being reformed.

()ne method of operation in the present invention consists in using thecatalyst successively for reforming and the first and second stagereactions. Thus naphtha is passed over the freshly regenerated catalystunder hydroforming conditions, e.g., 0.5 LHSV, 950 F., 200 p.s.i.g. Hrecycling 4 to 6 mols of H At the end of one-half to two hours, when thereforming activity has diminished because of mild coke lay-down, thenaphtha feed is cut out and replaced with cycle stock (first stage). Thetemperature may be raised to effect the desired reactions. Finally,after one to three hours on the first stage the reactor is switched tothe second stage by feedingaromatic oil from the first stage. At the endof this operation the catalyst is regenerated and readied for thehydroforming step.

In the initial stage of my process as the operating temperature israised, the dry gas, carbon make and con- Ma rate of about 2 liquidvolumes per hour per volume of catalyst: API/60" F. 21.8 Percent S Wt.percent-.. 2.38

5 Olefins wt. percent 26.8 Aromatics wt. percent.. 51.2 Br No. 20.9 N1.5251 B.P. Initial F 422 50% F 490 End-point F 568 At the end of fourone-hour runs, the products were separated and analyzed on an averagebasis as follows:

version to gasoline increase. Also the yield of recycle Product Wt men;oil decreases but its aromaticity increases. As shown by o't i eed thefollowing tabulation the initial stage aifords advantages over thethermal cracking of the light cycle stocks Liquid 79.4 conventionallyeffected. The gasoline produced in the gg ifi f-g reaction is higher inoctane number (e.g. 86 to 95 vs. 73) galrfbon on o t tuy st .7 and theby-product liquid is more valuable because of u a its low boiling rangeand viscosity and its utilization as 5 100-6 a feed stock fornaphthalene production.

Table I Feed (Light Cycle Oil):

API gravity, 60 F 21. 6 21.8 21.8 Percent S 1. 96 2. 38 2. 38 47.9 51.251.2 444 422 422 538 490 490 637 568 568 899 1,000 1,050 560 400 409 2 2H; mol ratio.. 1 5/1 Catalyst Thermal CrzOa/AlzOa CraOa/AlzO: Products(Wt. percent):

Dry gas 6. 4 6. 6 13. 6 aka 0. 6 1. 3 Gasoline (oi-400 F.) 29.1 29. a33.1 Research method octane no. neat (approx. 5 lbs. Reid 73 86 95Aromatic oil (400-600 F.) 57.7 45.6 API gravity at 60 F (17.8) (13. 7)Tar 64.5 3.5 2.9 API gravity at 60 F (4. l)

One form of the present invention can be described with reference to thedrawing which is a diagrammatic, simplified flow sheet of a systemcomprising separate reactors for the first and second stages,respectively, of

my process.

In the drawing, light cycle oil feed enters by way of line 1 intoreactor 2, which contains the hydroforming type catalyst. Ahydrogen-rich gas can enter reactor 2 by way of line 3 and the reactionproduct is conveyed in line 4 to separator 5. Gas is withdrawn fromseparator 5 through line 6. Aromatic gasoline is taken by way of line 7and the bottoms are removed in line 8. An aromatic-rich oil is drawnfrom separator 5 by way of line 9 and passed to the second stage reactor10 containing hydroforming catalyst. A hydrogen-rich gas enters reactor10 by way of line 11 and the reaction product passes through line 12 toseparator 13. Products withdrawn from separator 13 include aromaticgasoline, naphthalene, methyl naphthalene, other naphthalenics, andbottoms.

The following specific examples will serve to illustrate the presentinvention but they are not to be considered limiting.

EXAMPLE I A regenerated chromia-alumina catalyst (12% Cr O 2% MgO) wascharged to a reactor and brought to a temperature of about 1050 F. Acatalytic light cycle oil of the following analysis was charged to thereactor The liquid product was subjected to fractional distillation andproduced the following:

Product:

Gasoline (C to 400 F.) 33.4 wt. percent of Feed (0.037; sulfur, 58.1%aromatics, 9.7% olefins).

Aromatic oil (400-600 F.):

Wt. percent of feed PIT/60 F Wt. percent olefin 13.6. Wt. percentaromatics. 84.9.

1.575. Specific dispersion 269. Bottoms:

Wt. percent of feed 1.8.

1 Research method octane number of O to 400 F. gasoline was 95.8.

This yield includes 0 from gas.

An aromatic-rich oil stock produced according to the specific procedurenoted above but analyzing:

Boiling point F 400-600 API/60 F. 12.7 Specific dispersion 270 N 1.5740Wt. percent aromatics 81.3 Wt. percent olefins 13.2

was charged 'at a space velocity of 0.55-volumes per hour per volume ofcatalyst to a reactor containing a chr'ornia-alurnina catalyst (12% Cr O2% MgO). The catalyst was maintained at a temperature of about 1l09 The'liquid product was subjected to fractional distillation and produced thefollowing:

Product: and a p e of 400 P- -s and hydrogen Was P Gasoline (C4 e R) HWt, percent of Feed. plied to the reactor at a ratio of 5.7 mols per molof v (64.1% aromatics, 2.4% feed. The reaction continued for 2 hours andthe fol- Aromatic on 400 600F) 01mm)- lowing products were obtained, Wt.percent of Feed 4e. ARI/60 F 16. Wt. percent olefin 5.3. Wt. Percentpercent aromat 83.7. Product of Aromatic i l-5620 011 Feed Specificdispersion 258.

Bottoms:

Wt. percent of Feed 2.1. 8 1 Research method octane number of C t 5 o400 F. gasoline was 93.8. gag 6 This yield includes 0 from gas. Dry gas19. 3 Carbon on catalyst 4. 34 Sulfur on catalyst.

TotaL. 102-04 Wt. Percent Wt. Percent of Feed of 400-600 F. The liquidproduct was distilled and the following prod- Pmduc ucts were separated:Ngnh "1mm 3- 05 6'1 Alkylnaphthalenes 23. 2 47. 3 Wt. Percent Product ofAromatic Oil Feed 0 The aromatic-rich oil stock boiling 400600 F. was gg5 1 292 gig-g charged at a space velocity of about 1.0 volume per hourfl-methylnap 16.6 per volume of catalyst to a reactor containing amolyb- 3335 naphthalem 3g dena-alurnina catalyst 10.14% Moo -5% SiO Thecatalyst was maintained at about 985 F. and a pressure Research methodoctane number of C to 400 F. gasoline was equivaof of hydrogen andhydrogen was Supphed lent to iso;ootane +0.54 ml Eetraethyl lead. to thereactor at a ratio of 5.9 mols per mol of feed. ggl f g gjggf igfigf Thereaction was continued for one hour and the follow- The results of thetwo-stage process are summarized mg Products were Obtamed: below:

Stage I Stage II Combined Wt Percent Yields 1 Product of Aromatic OilFeed Run No 14 20 40 Temp. F. (400 lbs. p.s.i.g. H2)... 1,050 Liquid73,4 LHSV, approx 2 0 in g 1. 8 Prtzdufts (Wt. percent feed to Dry g14.0 S 91) I Carb on o tal st 10.9 Gasoline, C4 to 400 F. (V01. sulfu ion ca t aly s tn 0.2

percent) 33.40107) 21.0 46.9 (54. 3) Percent Arorn. (C to 400) 58.1 1001m 3 Percent S (C5 to 400) 0.03 R.M.O.N'. (C5 to 400) 95.8 106.5 98.7M.M.0.N. (0 to 400). 84. 106.5 90. 4 Heavier liquids fig% fi f g f::::ff 2f "iii The liquid product was distilled and the following prod-Methylnaphthalene. 16. 6 ucts were separated: Heavier oil 1. 8 6. 3 5. 4Carbon on catalyst. 1. 7 4. 3 4. 4 Dry gas 13. 6 19. 3 25. 6

Ultimate yields based on original light cycle oil feed and recycling F aalkylnaphthalenes to extinction in Stage II to produce maximum amountProduct 5 of naphthalene. All percentages noted above are calculated ona weight 1 69 basis unless otherwise specified.

EXAMPLE II aS(%lt]111e1(C4t0400F.) 19.5 a aene 22.7 A regeneratedmolybdena-alumina catalyst (10-1 aniistli lrra hthelene 15.7 M00 on Al O-5% SiO was charged to a reactor and gamma naphthalemc Pmductwa ottoms.5.8 brought to a temperature of about 950 F. The same feed employed inExample I was charged to the reactor at a rate of about 2 liquid volumesper hour per volume of catalyst. Hydrogen was supplied to the reactor atthe rate of 4.92 mols per mole of feed, under a pressure EXAMPLE III of400 p.s.i.g. At the end of 3 l-hour runs, the products were Separatedand analyzed on an average basis as 'As an example of severity controlin Stages I and II follows: of my process a number of operations wereconducted using the catalyst of Example I at a variety of tem- Pmductwtpercent peratures and space velocities. The light cycle oil feedOfFeed approximated that of Example I and the hydrogen was supplied tothe reactor at the rate of 7 mols per mol of feed and the total pressurewas 700 p.s.i.g. In these tests it was determined that at the reactionconditions 4.7 'v I sulfur on catalyst Lo gi en the followingtemperatures and space velocities could be employed; however, when usingthese conditions they should be selected so that the Stage II reaction 7ls'more severe than that of Stage I" as noted above.

Stage I, Ap- Stage II, Approximate proximate Temperature, T. Range ofSpace Maximum Velocity, Space Veloc- LHSV ity, LHSV 0. 2 to 0. 5 0. 10.25 to l. 0. 2

0. to 2. 0 0. 0. 9 to 4. 5 0. 7 1. 7 to 9. 5 1.2 3 to 20 2. 5

1 The minimum space velocity in Stage II could be about 0.1 LHSV at eachtemperature as a slower rate of feed would not be commercially easible.

I claim: 1

1. The method of producing naphthalene and high octane aromatic gasolinewhich comprises contacting in a first stage in the presence of hydrogena petroleum light cycle oil consisting essentially of aromatics andabout 40 to 65 percent of non-aromatic components with a hydroformingcatalyst having a non-combustible base at a temperature of about 900 to1200 F., a space velocity of about 0.1 to 20 LHSV and a pressure of atleast about atmospheric while efiecting conversion to an aromatic oilboiling above 400 F. and containing at least about 40 weight percent ofnaphthalenic aromatics and not more than about 30 weight percent ofnon-aromatics, said aromatic oil containing not more than about 10weight percent of naphthalene, contacting in a second stage under moresevere reaction conditions the aromatic oil with a hydroforming catalysthaving a non-combustible base, in the presence of hydrogen and at atemperature of about 900 to 1200 F., a space velocity of about 0.1 to 3LHSV and a pressure of at least about atmospheric to form naphthaleneand aromatic gasoline, and separating the gasoline and naphthalene.

2. The method of claim 1 in which the aromatic oil contains at leastabout 60 weight percent of naphthalenic aromatics and not more thanabout 20 weight percent of non-aromatics and the catalyst ischromia-alumina.

3. The method of producing naphthalene and high octane aromatic gasolinewhich comprises contacting in a first stage in the presence of hydrogena petroleum light cycle oil consisting essentially of aromatics andabout 40 to 65 percent of non-aromatic components with a hydroformingcatalyst having a non-combustible base at a temperature of about 900 to1200 F., a space velocity of about 0.1 to 20 LHSV and a pressure of atleast about atmospheric while efiecting conversion to high octanearomatic gasoline and an aromatic oil boiling above 400 F. andcontaining at least about 40 weight percent of naphthalenic aromaticsand not more than about 30 weight percent of non-aromatics, saidaromatic oil containing not more than about 10 weight percent ofnaphthalene, separating the high octane aromatic gasoline from thearomatic oil, contacting in a second stage under more severe reactionconditions the aromatic oil With a hydroforming catalyst having anon-combustible base, in the presence of hydrogen and at a temperatureof about 900 to 1200 F., a space velocity of about 0.1 to 3 LHSV and apressure of at least about atmospheric to form naphthalene and anadditional quantity of aro matic gasoline, and separating the gasolineand naphthalene.

4. The method of claim 3 in which the aromatic oil contains at leastabout 60 weight percent of naphthalenic aromatics and not more thanabout 20 weight percent of non-aromatics and the catalyst ischromia-alumina.

5. The method of claim 3 in which the catalyst employed in both reactionstages is chromia-alumina in which the severity in operating conditionsfor the first stage is defined by 1000 to 1100 F., 2 LHSV and 200 to 500p.s.i.g. hydrogen partial pressure and in which the severity inoperating conditions in the second stage is defined by 1050 to 1150 F.,0.5 LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure.

6. The method of claim 3 in which the catalyst employed in both reactionstages is molybdena-alumina, in which the severity in operatingconditions for the first stage is defined by 900 to 1000 F., 2 LHSV, andto 400 p.s.i.g. hydrogen partial pressure and in which the severity inoperating conditions in the second stage is defined by 900 to 1000 F.,1.0 LHSV and 200 to 500 p.s.i.g. hydrogen partial pressure.

7. The method of claim 3 in which beta-methylnaphthalene is alsoseparated from the product of the second stage and the remaining oilboiling above about 400 F. is recycled to the second stage reaction.

8. The method of claim 3 in which 2,6-dimethyl-naphthalene is alsoseparated from the product of the second stage, and the remaining oilboiling above about 400 F. is recycled to the second stage reaction.

9. The method of claim 3 in which the product of the second stageboiling above about 400 F. from which the naphthalene has beenrecovered, is recycled to the second stage reaction.

References Cited in the file of this patent UNITED STATES PATENTS2,328,828 Marschner Sept. 7, 1943 2,431,515 Shepardson Nov. 25, 19472,653,176 Beckberger Sept. 22, 1953 2,729,688 Anderson et al. Jan. 3,1956 2,758,062 Arundale et a1 Aug. 7, 1956 2,769,769 Tyson Nov. 6, 19562,780,661 Hemminger et a1. Feb. 5, 1957

1. THE METHOD OF PRODUCING NAPTHALENE AND HIGH OCTANE AROMATIC GSOLINEWHICH COMPRISES CONTACTING IN A FIRST STAGE IN THE PRESENCE OF HYDROGENA PETROLEUM LIGHT CYCLE OIL CONSISTING ESSENTIALLY OF AROMATICS ANDABOUT 40 TO 65 PERCENT OF NON-AROMATICS COMPONENTS WITH A HYDROFORMINGCATALYST HAVING A NON-COMBUSTIBLE BASE AT A TEMPERATURE OF ABOUT 900 TO1200*F., A SPACE WHEREBY OF ABOUT 0.1 TO 20 LHSV AND A PRESSURE OF ATLEAST ABOUT ATMOSPHERIC WHILE EFFECTING CONVERSION TO AN AROMATIC OILBOILING ABOVE 400*F. AND CONTAINING AT LEAST ABOUT 40 WEIGHT PERCENT OFNAPTHALENIC AROMATICS AND NOT MORE THAN ABOUT 30 WEIGHT PERCENT OFNON-AROMATICS, SAID AROMATIC OIL CONTAINING NOT MORE THAN ABOUT 10WEIGHT PERCENT OF NAPHTHALENE, CONTACTING IN A SECOND STAGE UNDER MORESEVERE REACTION CONDITIONS THE AROMATIC OIL WITH A HYDROFORMING CATALYSTHAVING A NON-COMBUSTIBLE BASE, IN THE PRESENCE OF HYDROGEN AND AT ATEMPERATURE OF ABOUT 900 TO 1200*F., A SPACE VELOCITY OF ABOUT 0.1 TO 3LHSV AND A PRESSURE OF AT LEAST ABOUT ATMOS-